Method for production of synthetic natural gas from crude oil

ABSTRACT

A process for the production of a pipeline gas of high BTU content from crude oil by hydrogasification of the crude oil. The crude oil is first vaporized in the presence of hydrogen and then gasified to form an effluent gas containing essentially methane, ethane, aromatic hydrocarbons, hyrogen and hydrogen sulfide. After separation of the aromatics and hydrogen sulfide from the effluent, the effluent is subjected to cryogenic separation of the hydrogen and a final catalytic conversion of the ethane to methane.

United States Patent Hegarty [111 I 3,870,481 Mar. 11,1975

METHOD FOR PRODUCTION OF SYNTHETIC NATURAL GAS FROM CRUDE OIL [76]Inventor: William P. Hegarty, 31 Fairway Ln.,

Wescoville, Pa. 18106 [22] Filed: Oct. 12, 1972 [21] Appl. No.: 297,012

[52] US. Cl 48/213, 48/197 R, 48/214 [51] Int. Cl C0lb 2/16 [58] Fieldof Search 48/211, 213, 214, 197 R; 260/449 M, 683.9; 208/362 [56]References Cited UNITED STATES PATENTS 3,183,181 5/1965 Rudbach 208/362X 3,363,024 1/1968 Majumdar et a1. 260/683.9

3,531,267 9/1970 Gould 48/213 3,537,977 11/1970 Smith 203/89 3,591,3567/1971 Thompson et a1. 48/213 OTHER PUBLICATIONS Kil1ingholme," C. J. P.De Winton, Gas Journal, Jan.

RESIDUAL OIL TO PARTIAL OXIDATION HYDROGEN FEED Primary Examiner-S.-Leon Bashore Assistant ExaminerPeter F. Kratz Attorney, Agent, orFirm-James C. Simmons; Barry Moyerman [57] ABSTRACT A process for theproduction of a pipeline gas of high BTU content from crude oil byhydrogasification of the crude oil. The crude oil is first vaporizedinthe presence of hydrogen and then gasified to form an effluent gascontaining essentially methane, ethane, aromatic hydrocarbons, hyrogenand hydrogen sulfide. After separation of the aromatics and hydrogensulfide from the effluent, the effluent is subjected to cryogenicseparation of the hydrogen and a final catalytic conversion of theethane to methane.

19 Claims, 4 Drawing Figures PRODUCT GAS LIQUID v AROMATICSWJFNWI'EUMARI 119.13 A 3.870.481

.SHEET 1 f 3 /0 /s 34 48 T v 22 4 OIL CRUDE on. Z oNE STAGEPURIFICATION-i, CRYOGENIC 7 TOPPING GASIFICATION SEPARATION r FUEL FOR mPLANT 2- 6 usE /3- /2 co 40 18M 56 vENT 24 v 54 5a OXYGEN HYDROGENSULFUR CATALYTIC Q PLANT PLANT 7 PLANT RICH PROD C 28 GAS SYN 5A;-REACTOR kaz PRODUCT GAS OIL F RESIDUAL OIL TO EESS PARTIAL OXIDATION 76LIQUID AROMATICS PTTTEF ITEUE- AR 1 1 1975 SHEET 2 or 3 STEAM /a'NAPHTHA FIRST CATALYTIC RICH GAS REACTOR METHANE AND ETHANE 50 /60 FROMHYDROGEN SEPARATION SECOND CATALYTIC RICHGAS REACTOR METHANATION UNITMETHOD FOR PRODUCTION OF SYNTHETIC NATURAL GAS FROM CRUDE OIL BACKGROUNDOF THE INVENTION 1. Field of the Invention.

This invention pertains to the production of synthetic natural gas(methane) by the gasification of crude oil. Crude oil is generallyseparated into high and low boiling point fractions and the low boilingpoint fractions are then treated at high temperature to produce aneffluent gas containing methane, ethane, acid gases such as hydrogensulfide, excess hydrogen, and residual aromatic constituents. Theeffluent from the gasification step is then subjected to furtherprocessing sequences where the acid gas and aromatic fractions areremoved and the hydrogen is separated, and finally the ethane is reactedto form additional methane and the methane or synthetic natural gas isfed to a pipeline for use in, among other things, residentialcommunities and industrial concerns.

2. Prior Art.

Treatment of crude oil or crude oil fractions to produce a syntheticpipeline gas either rich in hydrogen or rich in methane is shown in manyprocesses in the prior art.

One commercial process technique for producing synthetic natural gas byhydrocarbon gasification of naphtha distillate feed stocks is the streamreforming process operated either under severe or mild reformingconditions. In general, the steam reforming process is carried out atelevated temperatures at atmospheric or higher pressures in the presenceof a catalyst, generally a supported nickel catalyst. Hydrocarbons reactwith the steam to give a gas comprising hydrogen, carbon oxide, andmethane with the composition of the gas depending upon the conditions ofthe reaction. Steam reforming under severe conditions is disclosed inU.S. Pat. Nos. 3,103,423 and 3,063,936. British Pat. No. 981,726discloses a process for steam reforming under mild conditions wherein amethane rich gas is produced. This is illustrative of what is known inthe trade as the British Gas Council Catalytic Rich Gas Process (CRGProcess).

A second method of producing a synthetic pipeline gas, and one that isespecially rich in methane, is by the various thermal hydrocrackingtechniques. Thermal hydrocracking can be broken down into variouscategories. The first of these is the operation of the thermalhydrocracking unit without a coke bed or a catalyst; the second isoperation with a fluidized coke bed in order to control carbondeposition by periodic withdrawal of solids; and the last is theoperation with a hydroforming catalyst to influence the reaction.However, in the latter process the reaction remains largely thermalhydrocracking rather than catalytic hydrocracking at low pressures andhigh temperatures involved with steam being used to control cokedeposition (lay down) on the catalyst.

The thermal hydrocracking processes are exemplified in U.S. Pat. Nos.2,759,806, 2,882,138, 2,926,077, 3,124,436, 3,202,603, 3,484,219, and3,591,356. The foregoing Patents are illustrative of thermalhydrocracking processes that variously employ a fluidized bed,recirculation of the gas, or a multiple stage gasification system. Ingeneral, it is known that thermal hydrocracking processes orhydrocracking processes accomplished without the influence of a catalystwill produce large quantities of methane in the production of syntheticpipeline gases. Hydrocarbon feed stocks, boiling over a wide range, canbe used as the starting product for conventional thermal hydrocrackingprocesses. The conventional processes can operate at temperatures ofbetween 750 and 1650F. and at pressures from 800 to 3500 psig with largehydrogen requirements per barrel of liquid hydrocarbon feed. Highertemperatures are needed with the lighter feed stocks. Very heavy feedstocks may be processed at lower temperatures. Reactions include thermalcracking and hydrogenation, and production of a large volume of lightgases including methane is characteristic in large part because of therequired high temperature to produce cracking in the absence of acatalyst. In such processes, steam may be added to the reaction zone inorder to control the deposition of coke on the vessel interior.

Of particular interest is the process developed by the British GasCouncil and embodied in U.S. Pat. No. 3,363,024, which is called the GasRecycle Hydrogenator Process (GRH Process). This process has been usedsuccessfully in producing methane from light petroleum distillates,which can be subsequently processed to a synthetic natural gas forpipeline use. The process is basically a thermal hydrocracking one andis conducted at temperatures of between l290 and 1470F, and at pressureson the order of from to 1350 psig in the presence of a hydrogen gas. Thehydrogen gas reacts with the hydrocarbon feed in an exothermic reactionto produce gaseous hydrocarbons; mainly, methane and ethane. Productgases are formed from paraffins in the petroleum distillate and sidechains of aromatics in the distillate wherein the aromatic neuc'leus isunaffected. The aromatic constituents can then be separated from theproduct gases to produce a valuable benzol by-product. With thisprocess, the hydrogen gas can be either pure hydrogen or a gasconsisting predominately of hydrogen, such as the product gas from apartial oxidation unit or a steam reforming unit.

A third category of gasification processes which under certainconditions may be used in the process of the present invention includethe partial oxidation process. In these processes, a gaseous or liquidhydrocarbon feed is partially oxidized using oxygen, or air with lightfeeds, or steam with heavy feeds. Reactors are operated at pressures ofto 300 psig at temperatures of between 1800 and 3500F. Such reactorsproduce predominantly hydrogen and carbon monoxide and are well known inthe art. Two commercial processes that are offered for sale have beendeveloped by Shell and Texaco. The Shell process is shown in U.S. Pat.No. 2,971,829.

A fourth process for the manufacture of fuel gas is disclosed in U.S.Pat. No. 3,531,267 and combines a catalytic hydrocracking andgasification process.

Of the above processes, only those developed by the British Gas Counciland embodied in U.S. Pat. No. 3,363,024 and British Pat. No. 981,726 hasbeen found acceptable for producing a pipeline gas for use by the publicconsumer. The British Gas Council process has been found to beapplicable to feed streams with end points up to 365F (naphthas). It isnecessary to provide vaporization of the feed before introduction intothe gas recycle hydrogenator if carbon deposition is to be avoided. Onemethod is disclosed in U.S. Pat. No. 3,591,356 for utilizing the higherboiling feed stocks in order to convert the feed stock to methane,ethane, residual aromatics and hydrogen in the gas recycle hydrogenator,which product is used as a component of town gas. Another method isdisclosed in US. Pat. No. 3,124,436 wherein there is disclosed afluidized bed hydrogenation process. In each of these processes, carbonis formed in the gasifier.

The balance of the references cited above disclose processes that havenot achieved commercial acceptance for the direct manufacture ofsynthetic natural gas as of this time usually due to problems with theformation of carbon, coke, or tar in the reactor due to the reactions,as will be hereinafter explained.

SUMMARY OF THE INVENTION In accord with the present invention, there isprovided an improved process for producing a pipeline gas of highheating value from a crude oil. The process comprises basically thesteps of vaporizing a substantial portion of the crude oil at atemperature of between 600 and 1000F, thereafter introducing thevaporized crude oil and a hydrogenation gas into a gasification vesselmaintained at a temperature in excess of 1000F wherein the feed streamis gasified producing an effluent consisting essentially of hydrogen,hydrogen sulfide, methane, ethane, and residual aromatic hydrocarbons,thereafter cooling the effluent gas to room temperature and recoveringthe waste heat, drying the effluent and removing the hydrogen sulfideand residual aromatics from the effluent, cryogenically separating themethane and ethane from the hydrogen, and thereafter reacting the ethanewith steam to produce additional methane and carbon dioxide, andremoving the carbon dioxide. The two methane streams are combined anddischarged into a process pipeline or storage vessel. The methane fromthe gasifier does not have to be separated from the ethane during ethaneconversion. The process includes additional steps wherein the variousside streams can be used to produce sources of reactant or heat foroperating the plant utilities in order to achieve an efficient overallprocess. While it would be possible to use various hydrogasificationmethods, the gas recycle hydrogenator developed by the British GasCouncil is the preferred hydrogenation vessel for the method of theinstant application.

The hydrogen separated from the methane and ethane can be recycled tothe feed stream vaporization unit or the gasifier.

A key to the present invention is the vaporization step, which is alsodisclosed, comprising heating the oil and the hydrogenation gas,preferably hydrogen, and admixing the heated oil and gaseous hydrogen ina vaporizer and taking the vaporizer effluent and passing it in to thehydrogenation vessel, thereafter recovering the heat from the gasifiereffluent and using a portion of the gasifier effluent after cooling tocontrol the temperature of reaction of the gasification vessel therebypreventing unwanted coke formation in the gasification vessel.

Therefore, it is the primary object of this invention to provide animproved process for the production of synthetic natural gas from crudeoil.

It is another object of this invention to provide a method ofhydrogasifying crude oil.

It is a further object of this invention to provide a method for feedingheavier crude oil into a hydrogasification zone.

It is still a further object of this invention to provide a method ofproducing synthetic natural gas from crude oil wherein the by-productstreams from the main process stream are used in operating the overallprocess thereby providing process economies.

A BRIEF DESCRIPTION OF THE DRAWING FIG. 1 is a schematic drawing of theoverall process for producing a synthetic natural gas from a crude oil.

FIG. 1a is a schematic drawing of an alternative method of convertingnaphtha and ethane to methane and carbon dioxide.

FIG. 2 is a schematic diagram of the improved method of feeding thehydrogasification vessel in order to achieve an effluent that isconvertible to synthetic natural gas.

FIG. 3 is a schematic diagram of the test system used to verify thefeeding and hydrogasification steps of the present invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT There is shown in FIG.1 an overall process flow sheet for the method of the present invention.

In FIG. 1, numeral 10 indicates a feed stream preparation step in theoverall process wherein a crude oil stream 12 is subjected topreliminary separation wherein a naphtha fraction having 360F end pointis taken overhead through conduit 18 and used in a subsequent processingstep as will hereinafter be more fully explained.

The process stream 15 after preparation may be subjected to avaporization process as will be discussed in connection with FIG. 2,after which said stream comprises crude oil material having a +360Fboiling point and hydrogenation gas 13 which is injected into asingle-stage gasifier such as that disclosed in US. Pat. No. 3,363,024.The vaporizer and/or gasifier is represented in'FlG. 1 by block 16.

In the gasifier, about 83 percent of the oil is gasified at hightemperature to a mixture of methane, ethane, excess hydrogen, hydrogensulfide and residual aromatics. The 17 percent residue of the oil plus aportion of the aromatics separated from the gas after cooling are fed toa hydrogen plant 20. In the hydrogen plant 20, hydrogen (represented byarrow 13) is produced by the partial oxidation process for use in thegasifier or in vaporization of the feed stream in order to react withthe vaporized crude oil to formthe gasifier effluent 22. Oxygen for thepartial oxidation portion of the hydrogen plant 20 is supplied by aseparate oxygen plant 24, which plants are commercially available. Thebalance of the hydrogen plant 20 includes waste heat recovery systems,water gas shift, acid gas removal and methanation systems to producehigh-purity hydrogen by the reaction of CO and steam and subsequentremoval of hydrogen sulfide and C0 The CO from the hydrogen plant 20 isvented through conduit 26 and the hydrogen sulfide is passed bia aconduit 28 to a sulfur plant 30 wherein it is combined with the hydrogensulfide from the purification unit 34 and removed from the system by theclaus process as elemental sulfur as at 32. The gasifier effluent stream22 is cooled and passed through a purification section 34 wherein thegas is dried and the acid gases as well as benzene and other residualaromatics are removed.

Acid gases are taken to mean CO and hydrogen sulfide. The benzene andother residual aromatics are removed through conduit 36 and conductedthrough conduits 38 and 40 to the conduit 42, which receives theresidual oil fraction from the gasifier and the mixture introduced intothe hydrogen plant for production of hydrogen by the partial oxidationprocess. A portion of the benzene and residual aromatics can be used asa source of fuel for operating the plant. Hydrogen sulfide is removedthrough conduit 44 and sent to the sulfur plant 30. The effluent-frompurification step 34 in conduit 46 contains essentially hydrogen,methane and ethane and is sent to a cryogenic separation unit 48 whereinthe process stream is cooled, the methane and ethane are liquefied,separated from the hydrogen, vaporized, rewarmed compressed and removedthrough conduit 50, and the hydrogen through conduit 52. The hydrogen inconduit 52 is warmed to ambient and recycled to the gasifier or the feedstream preparation section 16 and respectively.

The combined methane-ethane liquid stream in conduit 50 is charged to acatalytic rich gas process unit 54 along with the naphtha that is takenfrom the feed stream preparation stage 10 via conduit 18. In thecatalytic rich gas section 54, the naphtha is desulfurized and thenaphtha and ethane are reacted with steam in an autothermic catalyticreaction to produce methane plus carbon dioxide. The carbon dioxide isseparated and vented through conduit 56 and the methane product inconduit 58 is dried and compressed to 1000 psig and discharged into aproduct pipeline or suitable storage receptacle.

In the catalytic rich gas unit 54, the methane produced in the gasifieris carried along with the process stream but does not enter into theoverall reaction. Alternatively, the methane from the gasifier can beseparated before the catalytic rich gas unit 54 thereby reducing theoverall size of unit 54. Another method of converting the naphtha ethaneis shown in FIG. 1a wherein the naphtha from conduit 18 is reacted withsteam 19 in a first catalytic rich gas reactor 54a thereby dischargingsteam, methane and carbon dioxide in conduit 160. The methane and ethanefrom the cryogenic separation unit 48 are injected in conduit 160through conduit 50 and the mixture is injected into a second catalyticrich gas reactor 5412 wherein the ethane and steam are reacted to formessentially methane and carbon dioxide. The effluent in conduit 162consists essentially of methane, carbon dioxide, carbon monoxide,

and hydrogen. This effluent is sent to a methanation unit whereinadditional methane is produced and the product gas is then sent to acarbon dioxide removal unit 166 through conduit 168. After carbondioxide removal, the effluent gas consisting of about 94% methane withhydrogen and a trace of carbon monoxide is sent into a product pipelinethrough conduit 170.

With a process as outlined in FIG. 1, a plant embodying such stages iscompletely self-sustaining regarding utilities. For example, by-productsteam is generated in the hydrogen plant, the sulfur plant, and thegasification plant. This steam is supplemented with an auxiliary steamgenerating plant to meet the requirements of the oxygen plant, thepurification process, cryogenic separation process and the catalyticrich gas section. Overall electrical power usage is neglibible sincecompressors and pumps are driven by condensing steam turbines. Fuel forauxiliary steam generators is supplied by the residual aromatics fromthe process, which are burned as a fuel. It is calculated that overallthermal efficiency (product gas heating value as a percent of feed oilheating value) is about 84 percent.

The overall process as set forth in connection wit FIG. 1 compriseseight separate steps each of which is within the perview of currenttechnology; however, such a combination has never before been shown orsuggested. However, the hydrogasification step 16 has not beensuccessful for heavier crude oil (higher boiling feed stocks).Hydrogasification is basically a highpressure, high-temperature,non-catalytic gas phase hydrocracking reaction. High molecular weighthydrocarbons are cracked and the fragments are saturated with hydrogen.Paraffins and naphthenes (cycloparaffins) are completely gasified tomethane and ethane in the hydrogasification reaction alkyl side chainsof the aromatics are gasified leaving residual benzene. Polycyclicaromatics are gasified leaving residual benzene. Polycyclic aromaticsare in part gasified also leaving residual refractory benzene and thesulfur is reacted to hydrogen sulfide. Overall, the reaction is veryhighly exothermic. The gas recycle hydrogenerator (GRH) as disclosed inUS. Pat. No. 3,363,024 is the only hydrogen gasification process provenin cornmercail operation. This process features an adiabatic back-mixedreactor that dealkylates feed aromatics and completely gasifiesparaffins and naphthenes (cycloparaffins) with a residence time ofbetween 10 and 50 seconds at temperatures in the order of 1290F to 1390Funder operating pressures of from to 1350 psig. The product gas fromsuch a gas recycle hydrogenator consists of methane, ethane, residualrefractory aromatics and excess hydrogen. The British Gas council haspublished its findings claiming that the only limitation on liquid feedstock is the ability to evaporate it as a gasifier feed. Commercialapplications have used naphthas with end points up to 365F. The GasCouncil has conductd pilot-plant tests with 635F end point gas oil withno gasifier carbon problems. However, in the latter case, pluggingproblems did develop and the Gas Council has accepted 650F as themaximum feeding end point for the process.

The Gas Council has identified three types of carbon depositions thatcan be found in the gas recycle hydrogenator. These are catalyticcarbon, which is deposited on catalytically active metal surfaces withattendant pitting corrosion of the metal surface; wall carbon, whichdeposits on a surface of low activity already coated with carbon; andspace carbon, which is produced as soot in the gas phase and exhaustedwith the gas product (effluent). Carbon formation is promoted byexcessively high operating temperatures, carbon monoxide and carbondioxide content and low gasifier hydrogen partial pressure. On the'otherhand, carbon formation is suppressed by sulfur and steam, with a highinternal recirculation rate to give a closer approach to completeback-mixing in uniform temperatures also tending to decrease carbonformation. At normal operating conditions, with negligible carbondioxide content, formation of space carbon and catalytic carbon with theaccompanying corrosion is eliminated by ten parts per million of sulfurin the feed distillate plus a small amount of steam. Sulfurconcentrations of to 300 parts per million in the feed may be requiredto eliminate wall carbon formation. Operating temperature and hydrogenpressure effects can be used in combination to avoid carbon formation,i.e. the normal 1380 operating temperature may be reduced to l320 whenhydrogen partial pressure is low. As the operating temperature islowered, however, the hydrogenation reaction rate decreases andparaffins and naphthenes may not be completely gasified. At temperaturesbelow 1290F. reaction rates decrease sharply and operation of thegasifier becomes unstable and the reaction may extinguish. 7 A

It has also been found thatthe hydrocarbon liquid/- hydrogen feed ratiomust be maintained within certain limits. At high ratios, hydrogenpartial pressure is low and the hydrogenation rate decreases withresulting cyclization of the reaction thereby causing a net productionof aromatics. In an extreme case, condensation to polycyclic aromaticscould result in tar and coke problems. Operation of the reactor includesmaintaining the temperature by regulating the feed preheat temperatures.At low ratios of hydrocarbon liquid to hydrogen feed less reactant ischarged, therefore, the heater reaction is less and the feed preheattemperatures increase. If such increases are excessive, the reactioncould initiate in the preheater with attendant difficulties therein.

The Gas Council has reported its efforts were directed to getting thehighest possible heating value gas direct from the gas recyclehydrogenator. Therefore, they sought to minimize the excess hydrogenthereby having a high ethane concentration in the effluent gas (about0.5 mole per mole of methane). They have not reported any investigationsusing excess hydrogen. One reported high-pressure run at I350 psia didhydrogenate ethane but the control was sluggish and operatingtemperature was unstable and carbon formation resulted in the vessels.If ethane hydrogenolysis is suddenly initiated with minimum excesshydrogemthere would be a sudden decrease in hydrogen partial pressureand the reacted temperature would rise. Such a combination would mostsurely produce carbon particularly when poor temperature control ispresent. Maintaining temperature control in the reactor with sufficientexcess hydrogen to suppress carbon formation the additional heat releasewould decrease the required preheat temperature for the 'products beingfed to the reactor.

Another major problem in oil hydrogasification is the formation of cokeand tar. Coke deposition could result in plugging of the reactor,thereby making the overall process uneconomical and inefficient. Tarcould cause fouling of heat exchange surfaces and prevent use of wasteheat boilers requiring quenching the gasifier product with consequentloss of potential hightemperature level heat recovery and decreasethermal efficiency. It is known that aromatics are precursors to tar andcoke formations. At elevated temperatures, aromatics condense to formheavy polycyclic molecules to give tars and ultimately coke. Maintaininghigh hydrogen partial pressure saturates olefinic cracking intermediatesand prevents them from cyclizing to produce aromatics and alsosuppresses the condensation of aromatics already present in the feed.Saturation of olefins by hydrogen also slows reactions because theolefms crack more easily than saturated compounds. Therefore, it followsthat in the presence of highpressure hydrogen, hydrocarbons can beheated to higher temperatures than normal without cracking or formationof tars or coke.

Recent advances in ethylene cracking technology have shown that aromaticcoke precursors that condense to coke in the laminar films on the hottube walls will form relatively fast in the bulk fluid at intermediatetemperature levels. The rate of coke precursor formation, relative tothe rate of the desired cracking reactions, was shown to decrease as thetemperature was raised. Accordingly, cracking furnaces were designed tominimize residence time in an intermediate temperature range. Cokeprecursor formation was sharply reduced; and with reduced precursors,tube wall temperatures at the outlet end were increased withoutincreased coking. The net result was development of the revolutionaryshort residence time ethylene cracking furnace. As an additionalbenefit, the tar formation was also reduced; and with lighter feedstocks, cracking oil transfer line waste heat boilers replacedquenching. This discovery leads to the conclusion that in the gasrecycle hydrogenator coke formation may be avoided even with heavy oilfeed stocks. The gas recycle hydrogenator is essentially a back-mixedreactor with the preheated feed (approximately -1000F) warmed togasifier effluent temperature (about 1380F) almost instantaneously. Inaddition, feed oil vapor concentrations are diluted to outletconcentrations immediately thereby minimizing self-condensationreactions; thus, formation of precursors at the intermediate temperatureand high feed concentration levels should be negligible and tar and cokeproduction minimized.

With this preliminary work in hand, it was discovered that the gasrecycle hydrogenator can accept feed stocks with higher boilingfractions when such feed stocks are gasified as shown in FIG. 2.

The basic gasification system is shown in FIG. 2 and consists of avaporizer 60, gasifier 62, a liquid aromatics separator 64, a productgas compressor 66, together with hydrogen and oil heaters 68 and 70respectively, and a gasifier product cooler 72. Such a system wouldoperate at 600 psig or higher with the crude oil feed 74 and thehydrogen feed 76, each being preheated and charged to the vaporizer 60.The oil 78 is further heatedby heater 80, which is submerged in the poolof liquid residual crude oil or heel as it is called in the trade. Thehydrogen after being preheated is sparged beneath the surface of theliquid residual crude oil 78 in vaporizer 60 so that the oil isvaporized into a mixture with the hydrogen and taken from the vaporizerthrough conduit 82 and conducted into the gasifier 62.

With whole crude oil, or naphtha topped crude oil, about percent byvolume of the oil would be evaporated with the residual oil beingwithdrawn through conduit 84 and used as feed for a partial oxidationplant to produce the required hydrogen. The level of the pool of liquidresidual oil and rate of residual oil withdrawal are determined bylevelcontrol 83. The vaporizer temperature would typically be in the range offrom 800 to 1000F depending upon the particular crude oil feed. Thehydrogen-oil vapors from the vaporizer in conduit 82 are injected intothe adiabatic hydrogasifier 62 and the exothermic gasification reactionoccurs. The oil is gasified to methane, ethane. and residual aromaticswith excess hydrogen and hydrogen sulfide which is generated by theconversion of sulfur in the oil. The entering feed jet 86 induces a highinternal recirculation rate in the gasifier along the path shown by thearrows. The contents of the entering jet 86 are essentially completelymixed and have substantially uniform composition and temperature. Thetemperature in the reactor is approximately 1400F and the averagereactor residence time is between 10 and seconds.

The hot effluent gases are removed through conduit 88 and cooled inwaste heat boiler 72 with heat recovery. After heat recovery, theeffluent is conducted to aromatic separator 64 through a conduit whereinthe condensed aromatic constituents are removed from the separatorthrough conduit 92 and are used for supplementary feed to the partialoxidation unit and process fuel. The resulting gas product from theseparator 64 is compressed in compressor 66 and pushed on to furtherprocessing. The gas at the exit end of the compressor consistsessentially of hydrogen, methane, ethane, hydrogen sulfide, anduncondensed aromatics. This product can be further processed inaccordance with the process described in FIG. 1. In FIG. 2, a portion ofthe cooled gas from the compressor 66 is taken through conduit 94 andrecycled through valve 96 into the reactor or hydrogasifier 62 toprovide cooling of the reactor to maintain the reacted temperature atthe desired temperature level of about 1400F. I

The foregoing process was verified in a test setup shown in FIG. 3 thatwas constructed and operated. This system consisted of a hydrogen supplyconduit 100, a hydrogen metering device (rotameter 102), electricallyheated hydrogen humidifier 104, an oil feed system 106 comprising an oiltank 108, and oil metering pump 110, a vaporizer 112 that waselectrically heated, a hydrogasification vessel 114, also electricallyheated, a product cooler-condenser 116, a liquid separator 118, a backpressure controller 120, product gas analysis system shown generally as122 consisting of a wet test meter 124 and a Ranerax meter 126, and atempering gas recycle system consisting of recycle gas compressor I28and recycle gas heater 130. In the above setup, electric heating wasprovided to the hydrogen humidifier, the vaporizer, the hydrogasifier,and the tempering recycle gas system and all high temperature transferlines were electrically heated to maintain the process temperature.

In operation, the gasification system pressure was maintained by theproduct gas back pressure controller 120. Hydrogasifier temperature wasautomatically controlled by the hydrogasifier electric trim heaters andthe hydrogen feed gas flow was manually set by hand control valve 132 tothe desired level as indicated by the rotameter 102. Hydrogen wassparged in beneath the surface of the water 134 maintained in thehumidifier and heated by heater 136 to becomesaturated with steam. Oilfrom tank 108 was pumped by the oil metering pump 110 into electricallyheated discharge tubing 138 mixed with the hydrogen and pumped into theelectrically heated vaporizer 112. The oil saturated hydrogen vaporsexited from the vaporizer 112 through conduit 140 and were introducedinto the hydrogasifier 114 through the nozzle 142. Residual unvaporizedoil was withdrawn from the vaporizer, as indicated by the level control144 in order to maintain a minimal level of liquid in the vaporizer, anddiscarded.

In view of the fact that overheating of the hydrogen oil hydrogasifierfeed vapors can cause coke depositions in the inlet nozzle 142 witheventual plugging, the feed inlet tube was jacketed with tempering gasflow, which entered through a pearlite insulated top section 148. Heatleakage from the hydrogasifier 114 was reduced by the pearlite and theheat was largely absorbed by the tempering gas avoiding excessiveheating of the feed vapor. The feed gas was injected into thehydrogasifier 114 at the top of the internal draft tube 150 at highvelocity. The entering jet induced a large internal recirculation rateto give an approach to complete back-mixing. Incoming oil vapors wereheated to the hydrogasifier temperature virtually instantaneously andreacted to give a product gas of methane, ethane, hydrogen sulfide,residual aromatics, and excess hydrogen in conduit 152. The testhydrogasifier was made of stainless steel reaction vessel approximately1.16 inches in diameter and 18% inches long. The axial draft tube 150was a 16-inch length of 20-gauge 96-inch stainless steel tubing. Theoverall volume of the test hydrogasifier was approximately 19.7 cubicinches.

The product gases exited the hydrogasifier 114 through an outlet tubeextending through the annulus from the bottom to the top of the drafttube. The effluent gases were cooled in a product cooler-condenser 116and passed through the liquid separator 118 where aromatic condensateliquid and water were collected and periodically drained through conduit154. The gas from the separator 118 was let down through a back pressurecontroller and with the rate measured in a wet test meter 124 andspecific gravity measured by a Ranerax instrument 126. Product gas wasvented to the atmosphere with periodic gas chromatographic analysisbeing made. Prior to pressure letdown, a portion of the gas wascompressed by the recycled gas compressor 128 heated to about 50 hotterthan the hydrogasifier feed gas temperature and passed as a temperinggas through the inlet feed gas jacketing. The tempering gas waswithdrawn through the hydrogasifier top and mixed with the hot gasifierproduct gas through conduit 156.

In an actual test run, Lagomedio crude oil was run in the abovetestsetup with the following results:

Gas composition, mole g 45.5 CH, 36.7 C 11 17.8-.

100.0 Aromatics Liquid Composition, wt

Benzene 36.7 Toluene 6.5 Cg aromatics 2.1 C -C aromatics 0.5 Naphthalene21.2 Anthrocene 5.4 Heavy ends 4.8 Residue & Loss 22.8

After 4 hours stable operation at these conditions, inspection of thereactor showed no significant coke deposition. The vaporizer showed nocoke deposition or tarry deposits except for a small quantity of friablecoke around the relatively cool (300F) inlet feed dip tube.

The test outlined above refers to a successful run feeding 76 weightpercent of a medium gravity (32.6 APl) lagomedio crude oil to thegasifier. Approximately 82 weight percent of the oil feed to thegasifier was gasified with 18 weight percent being recovered as aromaticcondensate; 76 weight percent vaporization is equivalent to about 1000FTPB cut point but even higher boiling consitutents were sure to be inthe gasifier feed. There were no problems experienced in the reactivefeed nozzle or vaporizer during this test.

In view of the foregoing, it was shown that the gas recycle hydrogenatorreactor system can be used with higher boiling feed stocks includingwhole crude oil; and it is also believed that residual oil fractions canbe used by direct partial vaporization at high pressure into hydrogen ofthe feed stock. Partial vaporization avoids the deposition and pluggingproblems that typically result when high boiling distillates or residualpetroleum fractions are evaporated to dryness. The residual liquid fromthe evaporation step with high boiling oil feed stocks provides asuitable feed to a partial oxidation process to supply the hydrogenrequirements for hydrogasification.

With high boiling distillates and residual oil feeds, high temperaturesare required for evaporation of the oil. Evaporation into hydrogen athigh pressure is beneficial in two respects. The high-pressure hydrogensuppresses cracking and coking reactions in the evaporator and alsodilutes the oil vapors reducing their partial pressure and decreasingthe temperature required to attain a given fraction vaporized.

It is within the scope of the present invention to use any of theconventional gasification schemes set forth in the art, however, the GasRecycle Hydrogenation is preferred.

If process economies dictated it, a portion of the carbon dioxideproduced in the process could be reacted with hydrogen in a methanationunit to produce athermal methane.

Having thus described my invention, what is desired to be secured byLetters Patent of the United States is set forth in the followingClaims.

I claim:

1. A method of producing a pipeline gas of high heating value from crudeoil comprising the steps of:

partially vaporizing a crude oil feed in the presence of hydrogen at atemperature of between 600 and 1000F, introducing the vaporized crudeoil feed and hydrogen into a gasification vessel maintained at atemperature in excess of 1000F wherein the feed stream is gasifiedproducing an effluent consisting essentially of hydrogen, hydrogensulfide, methane, ethane, and residual aromatic hydrocarbons;

cooling the effluent gases to room temperature and recovering waste heattherefrom to form water and aromatic condensate; and

removing the hydrogen sulfide, water and residual aromatics from saideffluent in a purification zone; cryogenically separating the methaneand ethane from the hydrogen in a purified effluent stream;

reacting steam with the ethane contained in the methane and ethanestream to produce methane and carbon dioxide; and

removing the carbon dioxide and discharging methane in a productreceiving device. 2. A method according to claim 1 wherein the crude oilfeed stream is subjected to an initial step wherein a naphtha fractionis removed.

3. A method according to claim 2 wherein the naphtha fraction is reactedwith steam in a catalytic process to produce methane and carbon dioxide,separating the carbon dioxide and combining the methane with methanefrom the ethane reaction.

4. A method according to claim 1 wherein there is included a hydrogengenerating plant which plant takes residual non-gasified crude oilbottoms resulting from the partial vaporization of crude oil feed,together with a portion of the aromatics from the gasifier effluent forproducing hydrogen for injection into the gasifier feed stream.

5. A method according to claim 1 wherein the hydrogen sulfide removedfrom the gasifier effluent is processed to elemental sulfur.

6. A method according to claim 1 wherein the gasification of the feedstream is carried out in an adiabatic hydrogasifier.

7. A method according to claim 1 wherein a portion of the product streamis cooled and after removal of residual aromatic and hydrogen sulfide isintroduced into the gasification vessel to maintain temperature controlin the gasifier vessel.

8. A method according to claim 1 wherein the cryogenically separatedhydrogen is warmed to ambient temperature and mixed with fresh hydrogenfor injection into the gasifier.

9. A method for producing a pipeline gas having a heating value of about1000 BTU/SCF from crude oil comprising the steps of:

subjecting a crude oil stream to a topping operation wherein there is a360F end point naphtha separation;

partially vaporizing the topped crude oil stream i the presence ofhydrogen at a temperature of between 600F and 1000F thereby admixing thecrude oil vapors and hydrogen to form a process stream;

gasifying the process stream to form an effluent consistin g essentiallyof methane, ethane, hydrogen sulfide, hydrogen and residual aromatichydrocarbons;

removing the residual aromatic hydrocarbons and hydrogen sulfide fromthe effluent in a purification unit; separating the hydrogen from themethane and ethane in the effluent;

reacting the naphtha from the crude oil topping steps and the methaneand ethane effluent with steam to produce a carbon dioxide and methaneeffluent; and

separating the carbon dioxide from the methane and introducing themethane into a product pipeline.

10. A method according to claim 9 wherein the hydrogen is separated fromthe methane and ethane efflu- 5 ent cryogenically and the hydrogen iswarmed to ambi ent and recycled to gasification.

11. A method according to claim 9 wherein the reaction of the naphthafraction and the methane and ethane effluent with steam is anautothermic catalytic reaction and the naphtha fraction is desulfurizedprior to said autothermic catalytic reaction.

12. A method according to claim 9 wherein the hydrogen sulfide separatedfrom the process stream effluent is treated to produce elemental sulfurby the Claus process.

13. A method according to claim 9 wherein the gasification of theprocess stream is effected in an adiabatic hydrogasifier wherein anexothermic reaction takes place to produce the gasifier effluent.

14. A method of gasifying a crude oil to produce a gaseous effluentconsisting essentially of hydrogen, methane, ethane, hydrogen sulfideand uncondensed aromatics comprising the steps of:

introducing a liquid crude oil into a vaporization vessel heated tobetween 800 and lOF;

maintaining a pool of liquid residual crude oil in said vaporizationvessel and introducing warmed gaseous hydrogen into the pool of liquidresidual crude oil thereby forming a mixture of vaporized hydrogen andcrude oil above the pool of liquid residual crude oil; and

withdrawing the vaporized crude oil hydrogen mixture and injecting saidmixture into a hydrogasification vessel wherein the mixture isrecirculated at a temperature of about 1400F to cause the mix ture toreact to form a gasifier effluent consisting essentially of hydrogen,methane, ethane, aromatic hydrocarbons and hydrogen sulfide and loweringthe temperature of said effluent and recovering heat thereby thuscondensing and separating condensible aromatic hydrocarbons.

15. A method according to claim 14 wherein a portion of the effluentstream is injected into the hydrogasification vessel to maintain thereaction temperature of said vessel.

16. A method according to claim 14 wherein a portion of the heat in saidgasifier effluent stream is recovered in a waste heat boiler beforeseparation of the condensible aromatic hydrocarbons.

17. A method according to claim 14 wherein residual oil from saidvaporization vessel is subjected to a partial oxidation process toproduce hydrogen for the vaporization step.

18. A method according to claim 14 wherein the crude oil is given a 360Fend point naphtha separation prior to being introduced into thevaporization vessel.

19. A method of gasifying a crude oil to produce a gaseous effluentconsisting essentially of hydrogen, methane ethane, hydrogen sulfide anduncondensed aromatics comprising the steps of:

injecting a crude oil stream into a vaporization vessel heated to atemperature of between 800 and 1000F and maintaining a pool of liquidresidual crude oil in said vaporization vessel;

sparging warmed gaseous hydrogen into said liquid residual crude oilbelow the surface of the pool thereby forming a mixture of hydrogen andvaporized crude oil above the pool;

withdrawing the vaporized hydrogen crude oil mixture and injecting saidmixture into a hydrogasification vessel;

allowing said mixture to circulate in said hydrogasification vessel toreact to produce a gasifier effluent consisting essentially of hydrogen,methane, ethane, aromatic hydrocarbons, and hydrogen sultide; and

lowering the temperature of said effluent and recovering heat therebythus condensing and separating condensable aromatic hydrocarbons.

1. A method of producing a pipeline gas of high heating value from crudeoil comprising the steps of: partially vaporizing a crude oil feed inthe presence of hydrogen at a temperature of between 600* and 1000*F,introducing the vaporized crude oil feed and hydrogen into agasification vessel maintained at a temperature in excess of 1000*Fwherein the feed stream is gasified producing an effluent consistingessentially of hydrogen, hydrogen sulfide, methane, ethane, and residualaromatic hydrocarbons; cooling the effluent gases to room temperatureand recovering waste heat therefrom to form water and aromaticcondensate; and removing the hydrogen sulfide, water and residualaromatics from said effluent in a purification zone; cryogenicallyseparating the methane and ethane from the hydrogen in a purifiedeffluent stream; reacting steam with the ethane contained in the methaneand ethane stream to produce methane and carbon dioxide; and removingthe carbon dioxide and discharging methane in a product receivingdevice.
 2. A method according to claim 1 wherein the crude oil feedstream is subjected to an initial step wherein a naphtha fraction isremoved.
 3. A method according to claim 2 wherein the naphtha fractionis reacted with steam in a catalytic process to produce methane andcarbon dioxide, separating the carbon dioxide and combining the methanewith methane from the ethane reaction.
 4. A method according to claim 1wherein there is included a hydrogen generating plant which plant takesresidual non-gasified crude oil bottoms resulting from the partialvaporization of crude oil feed, togEther with a portion of the aromaticsfrom the gasifier effluent for producing hydrogen for injection into thegasifier feed stream.
 5. A method according to claim 1 wherein thehydrogen sulfide removed from the gasifier effluent is processed toelemental sulfur.
 6. A method according to claim 1 wherein thegasification of the feed stream is carried out in an adiabatichydrogasifier.
 7. A method according to claim 1 wherein a portion of theproduct stream is cooled and after removal of residual aromatic andhydrogen sulfide is introduced into the gasification vessel to maintaintemperature control in the gasifier vessel.
 8. A method according toclaim 1 wherein the cryogenically separated hydrogen is warmed toambient temperature and mixed with fresh hydrogen for injection into thegasifier.
 9. A method for producing a pipeline gas having a heatingvalue of about 1000 BTU/SCF from crude oil comprising the steps of:subjecting a crude oil stream to a topping operation wherein there is a360*F end point naphtha separation; partially vaporizing the toppedcrude oil stream in the presence of hydrogen at a temperature of between600*F and 1000*F thereby admixing the crude oil vapors and hydrogen toform a process stream; gasifying the process stream to form an effluentconsisting essentially of methane, ethane, hydrogen sulfide, hydrogenand residual aromatic hydrocarbons; removing the residual aromatichydrocarbons and hydrogen sulfide from the effluent in a purificationunit; separating the hydrogen from the methane and ethane in theeffluent; reacting the naphtha from the crude oil topping steps and themethane and ethane effluent with steam to produce a carbon dioxide andmethane effluent; and separating the carbon dioxide from the methane andintroducing the methane into a product pipeline.
 10. A method accordingto claim 9 wherein the hydrogen is separated from the methane and ethaneeffluent cryogenically and the hydrogen is warmed to ambient andrecycled to gasification.
 11. A method according to claim 9 wherein thereaction of the naphtha fraction and the methane and ethane effluentwith steam is an autothermic catalytic reaction and the naphtha fractionis desulfurized prior to said autothermic catalytic reaction.
 12. Amethod according to claim 9 wherein the hydrogen sulfide separated fromthe process stream effluent is treated to produce elemental sulfur bythe Claus process.
 13. A method according to claim 9 wherein thegasification of the process stream is effected in an adiabatichydrogasifier wherein an exothermic reaction takes place to produce thegasifier effluent.
 14. A method of gasifying a crude oil to produce agaseous effluent consisting essentially of hydrogen, methane, ethane,hydrogen sulfide and uncondensed aromatics comprising the steps of:introducing a liquid crude oil into a vaporization vessel heated tobetween 800* and 1000*F; maintaining a pool of liquid residual crude oilin said vaporization vessel and introducing warmed gaseous hydrogen intothe pool of liquid residual crude oil thereby forming a mixture ofvaporized hydrogen and crude oil above the pool of liquid residual crudeoil; and withdrawing the vaporized crude oil hydrogen mixture andinjecting said mixture into a hydrogasification vessel wherein themixture is recirculated at a temperature of about 1400*F to cause themixture to react to form a gasifier effluent consisting essentially ofhydrogen, methane, ethane, aromatic hydrocarbons and hydrogen sulfideand lowering the temperature of said effluent and recovering heatthereby thus condensing and separating condensible aromatichydrocarbons.
 14. A METHOD OF GASIFYING A CRUDE OIL TO PRODUCE A GASEOUSEFFLUENT CONSISTING ESSENTIALLY OF HYDROGEN, METHANE, ETHANE, HYDROGENSULFIDE AND UNCONDENSED AROMATICS COMPRISING THE STEPS OF: INTRODUCING ALIQUID CRUDE OIL INTO A VAPORIZATION VESSEL HEATED TO BETWEEN 800* AND1000*F; MAINTAINING A POOL LIQUID RESIDUAL CRUDE OIL IN SAIDVAPORIZATION VESSEL AND INTRODUCING WARMED GASEOUS HYDROGEN INTO THEPOOL OF LIQUID RESIDUAL CRUDE OIL THEREBY FORMING A MIXTURE OF VAPORIZEDHYDROGEN AND CRUDE OIL ABOVE THE POOL OF LIQUID RESIDUAL CRUDE OIL; AND15. A method according to claim 14 wherein a portion of the effluentstream is injected into the hydrogasification vessel to maintain thereaction temperature of said vessel.
 16. A method according to claim 14wherein a portion of the heat in said gasifier effluent stream isrecovered in a waste heat boiler before separation of the condensiblearomatic hydrocarbons.
 17. A method according to claim 14 whereinresidual oil from said vaporization vessel is subjected to a partialoxidation process to produce hydrogen for the vaporization step.
 18. Amethod according to claim 14 wherein the crude oil is given a 360*F endpoint naphtha separation prior to being introduced into the vaporizationvessel.